Chloride assisted hydrometallurgical extraction of metal

ABSTRACT

A process for the extraction of zinc from a sulphide ore or concentrate containing copper and zinc includes subjecting the concentrate to pressure oxidation in the presence of oxygen and an acidic halide solution to obtain a resulting pressure oxidation slurry and subjecting the slurry to a liquid/solid separation step to produce a liquor containing copper and zinc in solution. The liquor containing the copper and zinc is subjected to a first solvent extraction with a copper extractant to remove copper from the solution and to produce a copper depleted raffinate. The copper depleted raffinate is subjected to a second solid extraction with a zinc extractant to produce a zinc depleted raffinate and the zinc depleted raffinate is recycled to the pressure oxidation step.

CROSS REFERENCE TO RELATED APPLICATIONS

This application is a continuation-in-part of U.S. patent applicationSer. No. 08/488,128 filed Jun. 7, 1995, now U.S. Pat. No. 5,650,057,which, in turn, is a continuation-in-part of U.S. patent applicationSer. No. 08/425,117 filed Apr. 21, 1995, now U.S. Pat. No. 5,645,708,which, in turn, is a continuation-in-part of U.S. patent applicationSer. No. 08/098,874 filed Jul. 29, 1993 which issued as U.S. Pat. No.5,431,788 on Jul. 11, 1995. The contents of the foregoing applicationsare incorporated herein by reference.

FIELD OF THE INVENTION

This invention relates to the hydrometallurgical treatment of metal oresor concentrates. In particular, it relates to the extraction of metalsfrom ores in the presence of halogen ions, such as chloride ions.

BACKGROUND OF THE INVENTION

Hydrometallurgical treatment of copper sulphide ores, such aschalcopyrite (CuFeS₂), is problematical because the severe conditionsrequired in a pressure oxidation step for the effective leaching ofcopper from these ores results in oxidation of the sulphide in the oreto sulphate, resulting in the generation of large amounts of acid whichrequires expensive neutralization. Attempts have been made to render thesulphide concentrate leachable under relatively milder conditions underwhich the sulphide would only be oxidized to elemental sulphur and notall the way through to sulphate. These attempts include the pretreatmentof the concentrate prior to the pressure oxidation step to render thesulphide concentrate more readily leachable, and the leaching of theconcentrate in the presence of chloride ions, such as described in U.S.Pat. No. 4,039,406. In this process, the copper values in theconcentrate are transformed into a solid basic copper sulphate fromwhich the copper values must then be subsequently recovered, asdescribed in U.S. Pat. No. 4,338,168. In the process described in U.S.Pat. No. 4,039,406 a significant amount (20-30%) of sulphide in the oreor concentrate is still oxidized to sulphate, resulting in greateroxygen demand during the pressure leach and the generation of sulphuricacid. This is particularly unfavourable for low grade concentrates,where the S/Cu ratio is high.

The present invention provides a process for the hydrometallurgicalextraction of copper and other metals in the presence of halogen ions,such as chloride and bromide in solution.

SUMMARY OF THE INVENTION

According to the invention there is provided a process for theextraction of copper from a sulphide copper ore or concentrate,comprising the steps of: subjecting the ore or concentrate to pressureoxidation in the presence of oxygen and an acidic halide solution toobtain a resulting pressure oxidation slurry and subjecting the slurryto a liquid/solid separation step to obtain a resulting pressureoxidation filtrate and a solid residue containing an insoluble basiccopper sulphate salt, characterized in that the pressure oxidation isconducted in the presence of a source of bisulphate or sulphate ionswhich is selected from the group consisting of sulphuric acid and ametal sulphate which hydrolyzes in the acidic solution and wherein theamount of the source of bisulphate or sulphate ions which is addedcontains at least the stoichiometric amount of sulphate or bisulphateions required to produce the basic copper sulphate salt less the amountof sulphate generated in situ in the pressure oxidation.

The process may further comprise the steps of recycling the pressureoxidation filtrate to the pressure oxidation; leaching the solid residuecontaining the basic copper sulphate salt in a second leaching with anacidic sulphate solution to dissolve the basic copper salt to produce aleach liquor containing copper sulphate in solution and a resultingsolid residue; separating the leach liquor from the solid residue;subjecting the leach liquor to a solvent extraction process to producecopper concentrate solution and a raffinate; and recycling the raffinateto the second leaching.

The pressure oxidation may be carried out at a predetermined molar ratioof H⁺ /Cu, where H⁺ represents the hydrogen ions in the acidic halidesolution and Cu represents the copper in the ore or concentrate, so thatthe pressure oxidation filtrate contains a first portion of the copperin the ore or concentrate and the basic copper salt contains a secondportion of the copper in the ore or concentrate and further comprisingthe steps of: separating the pressure oxidation filtrate and the basiccopper salt; leaching the basic copper salt in a second leaching stepwith an acidic sulphate solution to dissolve the copper salt to producea second copper solution and a solid residue; and subjecting thepressure oxidation filtrate and the second copper solution to solventextraction to obtain concentrated copper solution and copper depletedraffinate.

The pressure oxidation may be carried out at a temperature of from about115° C. to about 160° C., preferably about 150° C.

The process may further comprise the step of subjecting the pressureoxidation slurry to neutralization at a temperature above about 115° C.,preferably about 115° C. to 160° C., more preferably about 150° C.

Also according the invention there is provided a process for theextraction of copper from a sulphide copper ore or concentrate,comprising the steps of: leaching the ore or concentrate in a firstleaching step with an acidic chloride solution to produce a first coppersolution and an insoluble basic copper salt; separating the first coppersolution and the basic copper salt; leaching the basic copper salt in asecond leaching step with an acidic sulphate solution to dissolve thecopper salt to produce a second copper solution and a solid residue; andsubjecting the first and second copper solutions to solvent extractionwith an organic extractant to produce concentrated copper solution forelectrowinning of copper therefrom.

Further according to the invention there is provided a process for theextraction of copper from a sulphide copper ore or concentrate,comprising the steps of: subjecting the concentrate to pressureoxidation in the presence of oxygen and an acidic halide solution toobtain a resulting pressure oxidation slurry; and subjecting the slurryto neutralization at a temperature above about 115° C.

Also according to the invention there is provided a process for theextraction of zinc from a sulphide ore or concentrate containing copperand zinc, comprising the steps of: subjecting the concentrate topressure oxidation in the presence of oxygen and an acidic halidesolution to produce a liquor containing copper and zinc in solution;subjecting the liquor to a first solvent extraction with a copperextractant to remove copper from the solution and to produce a copperdepleted raffinate; subjecting the copper depleted raffinate to a secondsolvent extraction with a zinc extractant to produce a zinc depletedraffinate; and recycling the zinc depleted raffinate to said pressureoxidation.

Further according to the invention there is provided a process for theextraction of precious metals from a copper sulphide ore or concentrateby treating a leach residue of said ore, comprising the steps of:removing elemental sulphur from said leach residue to obtain a lowsulphur residue; and subjecting the low sulphur residue to an oxidativeleach at elevated temperature and pressure to oxidize sulphur andprecious metal compounds present in the low sulphur residue to produce aliquid containing the precious metals in solution.

Also according to the invention there is provided a process for theextraction of precious metals from a copper sulphide ore or concentrate,comprising the steps of: subjecting the concentrate to pressureoxidation in the presence of oxygen and an acidic halide solution toobtain a resulting pressure oxidation slurry; and flashing the slurry toatmospheric pressure in a two-stage let-down in which the first stage isat a temperature of about the freezing point of elemental sulphur.

BRIEF DESCRIPTION OF THE DRAWINGS

The invention will now be described by way of examples with reference tothe accompanying drawings, in which:

FIG. 1 is a flow diagram of a hydrometallurgical copper extractionprocess according to one embodiment of the invention, which is suitablefor the treatment of high grade copper ores or concentrates;

FIG. 2 is a flow diagram of a hydrometallurgical copper extractionprocess according to another embodiment of the invention, which issuitable for the treatment of medium and lower grade copper ores orconcentrates;

FIG. 3 is a schematical illustration of a pressure vessel to illustrateanother embodiment of the process for the treatment of concentrates witha high sulphur to metal ratio;

FIG. 4 is a flow diagram of a hydrometallurgical process for theextraction of metals from a copper-zinc sulphide concentrate accordingto another embodiment of the invention; and

FIGS. 5A and B show a flow diagram of a further embodiment of theprocess according to the invention for the recovery of precious metalsfrom an ore or concentrate.

DETAILED DESCRIPTION OF PREFERRED EMBODIMENTS

The different embodiments of the process according to the invention maybe used to treat a range of copper and other metal concentrates in whichthe grade of copper varies from low, i,e. about 15% copper or less, tohigh grade, i.e. about 35% copper or more.

Broadly, the process comprises a pressure oxidation step which takesplace in the presence of oxygen and an acidic solution of halide ions,e.g. chloride or bromide, and sulphate ions. More specifically, theprocess further includes an atmospheric leach stage, one or more solventextraction stages and an electrowinning stage. Different grades ofconcentrate require different treatment in the pressure oxidation stage,requiring different modes of operation. These modes of operation aretermed Mode A, Mode B and Mode C, respectively. In Mode A, which iseffective when high grade copper ores are leached, copper is not leachedin the pressure oxidation stage. In Modes B and C, which are effectivewhen medium and low grade copper ores are leached, copper is leached inthe pressure oxidation stage.

Each of the modes of operation will now be described in turn.

Process Mode A

FIG. 1 is a flow diagram of Mode A. The process comprises a pressureoxidation stage 12 in a pressure oxidation vessel or autoclave, anatmospheric leach stage 14, primary and secondary solvent extractantstages 16 and 18, respectively, and an electrowinning stage 20.

In the pressure oxidation stage 12, all copper minerals are converted tobasic copper sulphate, CuSO₄.2Cu(OH)₂. The treatment is carried out withoxygen in the presence of an acidic chloride solution. Oxygen, as wellas HCl and H₂ SO₄ are introduced into the autoclave for this purpose.The temperature in the autoclave is about 130°-150° C. and the pressureabout 100-200 psig (about 800 kPa to about 1500 kPa). This is totalpressure comprising oxygen pressure plus steam pressure. The retentiontime is about 0.5-2.5 hours and the process is normally carried out in acontinuous fashion in the autoclave. However, the process can also becarried out in a batch-wise fashion, if desired.

The solids content in the autoclave is maintained at about 12-25%, i.e.150-300 g/L solids as determined by the heat balance and viscositylimitations.

The slurry produced in the autoclave is discharged through a series ofone or more flash tanks 22 to reduce the pressure to atmosphericpressure and the temperature to 90°-100° C. The liquid part of theslurry is referred to as the product solution from the pressureoxidation stage 12 and is indicated by reference numeral 21.

The slurry from the flash tank(s) 22 is filtered, as shown at 24, andthe resultant filter cake is washed thoroughly to remove entrainedliquor as much as possible.

The pressure oxidation filtrate from the filtration 24 is recycled tothe pressure oxidation stage 12 but there is a small bleed of about 5%,as shown at 26. This bleed 26 is determined by the concentration of thesoluble metals in the ore or concentrate which may dissolve during thepressure oxidation stage 12. The bleed 26 is treated at 28 with lime toremove metals such as zinc and magnesium as solid residues, which arepresent in the copper concentrate, and to counteract buildup of thesemetals in the pressure oxidation circuit. The pressure oxidation circuitis the circuit from the pressure oxidation stage 12 to the flash tank(s)22 to the filtration 24 to the bleed 26 and back to the pressureoxidation stage 12. It is indicated by reference numeral 23.

The bleed 26 is subject to a solvent extraction, as shown at 27, priorto the bleed treatment 28. The solvent extraction 27 is carried out bymeans of a suitable organic extractant to remove copper from the bleed26. This solvent extraction is associated with the solvent extractionstages 16 and 18 and will be referred to again when the latter twosolvent extraction stages are described.

Prior to the pressure oxidation stage 12, the copper concentrate isfirst subjected to a regrind, as shown at 30, to reduce the particlesize to about 97% minus 325 mesh, which corresponds to P80 (80% passing)15 micron. The regrind 30 is carried out in solution recycled from thebleed treatment 28. Thus, the slurry from the bleed treatment 28 issubjected to a liquid/solid separation, as shown at 32, and the solutionis recycled to the regrind 30 and the zinc/magnesium bleed residue isdiscarded, as shown at 17.

The solution which is recycled to the regrind 30 is an alkaline chlorideliquor at about pH 10. Use of this liquor minimizes water input into thepressure oxidation circuit 23 which is important in maintaining heatbalance and in preserving the chloride solution in the pressureoxidation circuit 23 as much as possible.

As stated above, copper is not leached in the pressure oxidation stage12 but is converted to an insoluble basic copper salt. The feed solutionto the pressure oxidation stage 12, which is the leach liquor beingrecycled from the filtration 24 is indicated by reference numeral 25.Although there is copper present in the feed solution 25, there is noadditional copper leached, i.e. the process is operated so that thecopper concentration in the feed solution 25 to the pressure oxidationstage 12 is equal to the copper concentration in the product solution 21from the pressure oxidation stage 12. This is indicated as: Δ Cu²⁺ !=0.

The feed solution 25 to the pressure oxidation stage 12 contains about15 g/L Cu and 12 g/L Cl, together with about 30-55 g/L sulphuric acid.The acid is added in the form of make up H₂ SO₄ (usually 93%). Theproduct solution 21 from the pressure oxidation stage 12 also containsabout 15 g/L Cu and 11-12 g/L Cl but is at about pH 3. There issubstantially no acid left in the product solution 21 as it is allconsumed in the pressure oxidation stage 12 to form the basic coppersalt.

As referred to above, the liquid feed 25 to the pressure oxidation stage12 is made up partly of recycled filtrate to which H₂ SO₄ is added. Theimmediate effect of adding the acid to the filtrate is to increase theacidity of the filtrate which is fed to the autoclave for the pressureleaching stage 12, but the most important effect, surprisingly, has beenfound to be that the addition of the acid, or more specifically thesulphate ions, actually suppresses the oxidation of sulphur emanatingfrom the concentrate in the pressure oxidation stage 12.

Typically the oxidation of sulphur that is experienced if no acid isadded is about 25-30% of the feed sulphur in the concentrate, as is thecase with the process described in U.S. Pat. No. 4,039,406. However, ifacid is added, it has been found that the sulphur oxidation to sulphateis reduced to about 5-10%. This improvement has substantial beneficialeffects on the hydrometallurgical extraction process. The oxidation ofsulphur to sulphate creates additional costs in several ways, such asadditional oxygen required for the reaction, additional reagent requiredto neutralize the acid so formed by the oxidation and provision must bemade for heat removal due to the oxidation of sulphur to sulphate whichis very exothermic. This actually limits the throughput of the autoclavein which the pressure leaching stage 12 takes place.

The chemistry of the reaction in the pressure oxidation stage 12 isbelieved to be altered by the addition of the acid as follows:

No acid addition: ##STR1## With acid addition: ##STR2##

In both reactions, the copper is precipitated in the form of a basiccopper salt, which has been found to comprise mostly basic coppersulphate.

In the first reaction it appears that the sulphate of the basic coppersulphate is supplied by oxidation of the feed sulphur in theconcentrate, whereas in the second reaction it appears to be supplied bythe sulphate ions in the acid which is added to the autoclave, thusobviating the need for the oxidation of sulphur to sulphate. Thus, inthe second reaction, there is a net consumption of sulphate ions to formthe basic copper salt. The amount of sulphuric acid needed to suppresssulphur oxidation has been found experimentally to be about 25 to 75grams per liter, depending on the type of concentrate and the percentagesolids in the concentrate.

In actual test work, there is more sulphur oxidation than is predictedby either reaction. The first reaction predicts one sixth or 16.7% ofthe sulphur to be oxidized, whereas experimentally about 25%-30% isfound. With acid addition, experiments indicate that about 2-16% sulphuris oxidized to sulphate, rather than the zero oxidation that would bepredicted if the second reaction as written was the only reaction takingplace. Therefore, these reaction equations do not reflect exactly whatis happening in the pressure leaching stage 12 but are only anapproximation.

Chloride is conserved as much as possible in the pressure oxidationcircuit 23 but typically about 3-10% chloride is lost per pass into thesolid product at the filtration 24. Thus, the chloride must be made upby the addition of HCl or another source of chloride to provide 12 g/Lchloride in the feed solution 25. The chloride losses are minimized bythorough washing of the solids from the pressure oxidation stage 12 onthe filter 24. The amount of wash water is constrained by therequirement to maintain a water balance in the pressure oxidationcircuit 23. The only water loss from the circuit 23 is in the steam 9from the flashing step 22 and in the filter cake after the filtration24. Hence, the need to use the recycled solution from the bleedtreatment 28 to slurry up the concentrate in the grinding step 30, andthus minimize fresh water input from the concentrate to the pressureoxidation step 12.

It has been found to be advantageous to maintain at least 15 g/L Cu inthe product solution 21 from the pressure oxidation stage 12 so as tocounteract chloride loss in the form of solid basic copper chloride,CuCl₂.3Cu(OH)₂, which can occur if insufficient copper is present insolution to allow basic copper sulphate to form:

    4CuCl.sub.2 +6H.sub.2 O→CuCl.sub.2. 3Cu(OH).sub.2 +6HCl(3)

This reaction can be counteracted by the addition of sufficient acidinto the autoclave during the pressure oxidation stage 12 to maintain atleast enough copper in solution to satisfy the stoichiometricrequirements for Cl as CuCl₂. For 12 g/L Cl in solution, thestoichiometric amount of Cu is: ##EQU1##

Thus, 15 g/L Cu is a safe minimum to prevent a significant chloride lossin the form of the basic copper salt.

On the other hand, the copper concentration in the product solution 21from the pressure oxidation stage 12 should be kept as low as possibleto counteract the formation of CuS by the reaction of elemental sulphurwith aqueous copper sulphate. This reaction can occur during thepressure oxidation stage 12 or in the slurry after discharge from theautoclave but before the filtration step 24:

    3CuSO.sub.4 (aq)+4S.sup.0 +4H.sub.2 O→3CuS(s)+4H.sub.2 SO.sub.4(4)

This reaction is particularly undesirable because CuS is insoluble inthe dilute acid conditions of the atmospheric leaching stage 14. Thus,the copper is not recovered and this results in the loss of copper tothe final residue.

To counteract the formation of CuS it is necessary to keep the copperconcentration in the product solution 21 as low as possible, i.e. below30 g/L for some concentrates. The tendency to CuS formation isapparently related to the type of concentrate being treated, with themedium to high grade concentrates being more susceptible to CuSformation. Thus, although a high copper concentration in the productsolution 21 does not present a problem with the low grade concentrates,it cannot be tolerated with the higher grade concentrates.

As is known to date, high grade concentrates, i.e. above 35% copper, arebest treated to produce as low a copper concentration in the productsolution 21 as possible, i.e. below 25 g/L Cu.

Given the need to maintain at least 15 g/L Cu in solution in thepressure oxidation circuit 23, there is an optimum range of copperconcentration of from 15 to 25 g/L Cu for high grade concentrates. Withmedium grade concentrates, the upper limit can be stretched considerablyand for low grade ore, the copper concentration does not play asignificant role.

The copper concentration in the pressure oxidation filtrate 29 can becontrolled simply by adding the required amount of acid into the feedsolution 25 to the pressure oxidation stage 12. More acid results in ahigher copper concentration due to the dissolution of the basic coppersulphate:

    CuSO.sub.4.2Cu(OH).sub.2 (s)+2H.sub.2 SO.sub.4 →3CuSO.sub.4 (aq)+4H.sub.2 O                                           (5)

The addition of about 1 g/L acid results in an increase in copperconcentration of about 1 g/L. The actual concentration of acid requiredis determined empirically by comparing the assays of feed solution 25 tothe pressure oxidation stage 12 and the product solution 21 from thepressure oxidation stage 12 to satisfy Δ Cu² +!=0. The volume ofsolution in the circuit 23, however, is determined by the heat balance.

The percentage by weight of solids in the feed of copper concentrateslurry to the pressure oxidation stage 12 can be varied at will. Theweight of concentrate solid fed to the pressure oxidation stage 12 isdetermined by the amount of copper to be recovered. The weight of thesolution is determined mainly by the heat balance in the pressureoxidation stage 12.

The desired operating temperature in the pressure oxidation stage 12 isabout 150° C. and the heat must be supplied largely by the heat ofreaction of the sulphide minerals with the high pressure oxygen in theautoclave. For high grade concentrates, such as will be treated by theProcess Mode A currently being described, this means a relatively lowS/Cu ratio and thus a smaller heat production per tonne of coppertreated in the autoclave. Much of the heat evolved is due to oxidation,not of copper, but of the other two main elements in the concentrate,iron and sulphur. If the grade of the concentrate is high, then theratio of S/Cu and Fe/Cu is low, hence a lower heat production.

To reach operating temperature from a starting temperature of say 50° to80C., which is typical for the pressure oxidation filtrate 29 which isrecycled after the filtration 24, it is necessary to control the amountof water that must be heated, since this is the main heat sink in thepressure oxidation stage 12. It is impractical to cool or heat theslurry inside the autoclave by indirect means, such as by means ofheating or cooling coils, because of rapid scale formation on allsurfaces, particularly heat exchangers, leading to very poor heattransfer characteristics. Direct heating or cooling by injection ofsteam or water is also impractical due to water balance considerations.Therefore, it is required that the heat balance be maintained bybalancing heat production from reaction heat with the heat capacity ofthe feed materials, i.e. the feed solution 25 being recycled and theconcentrate slurry. The main variable that can be controlled here is thevolume of the feed solution 25. This is one of the distinguishingfeatures between Modes A and B. In Process Mode B, still to bedescribed, where greater sulphur oxidation is experienced, the heatevolution is much greater, expressed as heat per tonne of copperproduct. Therefore, it is possible to use more solution volume in thefeed 25 to the pressure oxidation stage 12.

Once the solution volume is fixed, the acidity of the solution can bedetermined, since the total mass of acid is determined by the need tomaintain Δ Cu² +!=0. Typically, for a high grade concentrate, about35-55 g/L acid will be required.

It has been found to be beneficial to add small concentrations ofcertain surfactants which change the physical and chemicalcharacteristics of liquid elemental sulphur (S⁰) in the autoclave duringthe pressure oxidation stage 12. Surfactants such as lignin sulphonateand quebracho added to the pressure oxidation feed solution 25 in smallamounts, i.e. 0.1 to 3 g/L can reduce the viscosity of the liquidsulphur and also change the chemistry in the autoclave.

Additions of surfactants can reduce sulphur oxidation in ways that arenot well understood, but are beneficial to the process. It is believedthat this is due to lower viscosity, resulting in lowered tendency forliquid sulphur and solids to be held up within the autoclave, thusreducing the retention time for these materials, and hence the reducedtendency for sulphur oxidation to occur.

Also it has been found that more complete reaction of the copperminerals takes place if surfactants are added, apparently because oflower viscosity sulphur, which does not "wet" unreacted sulphideminerals, and thus allows the desired reaction to proceed to completion.

Reaction (5) describes how adding sulphuric acid to the pressureoxidation feed 25 will control the copper concentration in the pressureoxidation filtrate 29. The overall reaction for the pressure oxidationwith sulphuric acid addition for a chalcopyrite ore is given by reaction(2) above.

A similar reaction can be written using CuSO₄ as the source of sulphideions instead of H₂ SO₄ :

    3CuFeS.sub.2 +15/4O.sub.2 +3H.sub.2 O+3/2CuSO.sub.4 →3/2CuSO.sub.4.2Cu(OH).sub.2 +3/2Fe.sub.2 O.sub.3 +6S.sup.0(6)

It is noteworthy that there are 3/2 moles of sulphate required as coppersulphate in reaction (6) compared to one mole of sulphuric acid inreaction (2). Therefore, if CuSO₄ is to be used as the source ofsulphate ions instead of sulphuric acid, it is necessary to use 1.5times as many moles of CuSO₄. To take this into account, the inventorhas developed the concept of Excess Sulphate Equivalent, which allowsthe calculation of how much acid to add to the pressure oxidation feedsolution 25 in order to achieve a target copper concentration and stilltake into account reaction (6).

By taking reaction (6) into account, it is possible to calculate "apriori" the amount of acid required for constant copper concentration inthe pressure oxidation filtrate 29. The concept of Excess SulphateEquivalent is helpful:

Excess Sulphate Equivalent is equal to the sulphate available in thepressure oxidation feed solution 25 for formation of basic coppersulphate during the pressure oxidation stage 12. The sulphate availableis that which is in excess of a defined Base Level of CuSO₄ and CuCl₂.

Base Level of CuSO₄ and CuCl₂ is sufficient to support chloride insolution at 12 g/L in the form of CuCl₂ and, in addition, about 4.3 g/LCu as CuSO₄. The concentration of CuCl₂ corresponding to 12 g/L chloridein solution is 134.5/71*12=22.7 g/L CuCl₂, which contains 10.7 g/L Cu insolution. The additional 4.3 g/L copper therefore means a total of 15g/L copper combined as CuCl₂ and CuSO₄ in the Base Level.

Sulphate available is then the total sulphate as CuSO₄ less the BaseLevel. For instance, if the total copper concentration is 28 g/L in thepressure oxidation filtrate 29, then the sulphate available is 28-15=13g/L Cu*98/63.5=20 g/L H₂ SO₄ as available sulphate from CuSO₄.

Excess Sulphate Equivalent (ESE) is then calculated from the availablesulphate from CuSO₄ by dividing by 1.5:

    ESE={Available Sulphate as CUSO.sub.4 }/1.5

Thus, in the example of 28 g/L total copper concentration or 20 g/Lavailable sulphate from CUSO₄, there is 20/1.5=13.3 g/L ESE from CuSO₄.

Finally, if the target free acid equivalent is, say, 52 g/L H₂ SO₄ inthe pressure oxidation feed solution 25, then the amount of acidrequired is 52 less the ESE (13.3 g/L) or 38.7 g/L H₂ SO₄. This is theamount that must be added to the feed solution 25 to the pressureoxidation stage 12 to produce a constant copper concentration in thepressure oxidation filtrate 29, i.e. the Base Level of 15 g/L Cu.

Other reactions can be written using Fe₂ (SO₄)₃ and ZnSO₄ as the sourceof sulphate ions instead of H₂ SO₄. In the case of ZnSO₄, the zinc isassumed to hydrolyze to basic zinc sulphate, ZnSO₄.3Zn(OH)₂, which is abasic salt of Zn analogous to basic copper sulphate. These reactions aregiven below as reactions (7) and (8).

    3CuFeS.sub.2 +15/4O.sub.2 +2H.sub.2 O+1/3Fe.sub.2 (SO.sub.4).sub.3 →CuSO.sub.4.2Cu(OH).sub.2 +11/6Fe.sub.2 O.sub.3 +6S.sup.0(7)

    3CuFeS.sub.2 +15/402+13/3H.sub.2 O+4/3ZnSO.sub.4 →CuSO.sub.4.2Cu(OH).sub.2 +6S.sup.0 +Fe.sub.2 O.sub.3 +1/3{ZnSO.sub.4.3Zn(OH).sub.2.4H.sub.2 0}                 (8)

The solids from the pressure oxidation stage 12 after the filtration 24,are treated in the atmospheric leaching stage 14 at about pH 1.5 to pH2.0 using raffinate from the primary solvent extraction stage 16, whichis acidic, to dissolve the basic copper sulphate. The leaching 14 takesplace at a temperature of about 40° C. for a retention time of about15-60 minutes. The percentage solids is typically about 5-15% or about50-170 g/L, although it is possible to operate the process outside thisrange.

During the atmospheric leaching stage 14, the basic copper saltsdissolve almost completely with very little of the iron present in theconcentrate going into solution.

Typically, the leach liquor 33 produced after the liquid/solidseparation 34 contains about 10-20 grams per liter copper, depending onthe percentage solids feed to the leach 14, with 0.1-1.0 g/L iron andabout 0.1-1.0 g/L chloride. Much of this iron and chloride are derivedfrom the feed raffinate 37 rather than the solids from pressureoxidation, i.e. they are recycled. Typically about 0.1-0.2 g/L iron andchloride dissolve per pass.

The copper extraction has been found to be about 95-98% based on theoriginal feed to the pressure leaching stage 12. Iron extraction tosolution has been found to be less than about 1%.

The slurry 31 from the atmospheric leaching stage 14 is difficult if notimpossible to filter, but settles well. In view of the need to wash theleach solids very thoroughly, the slurry 31 is therefore pumped to acounter current decantation (CCD) wash circuit, symbolically indicatedas a solid/liquid separation 34 in FIG. 1. In the CCD circuit 34, thesolids are fed through a series of thickeners with wash water added inthe opposite direction. By this method, the solids are washed andentrained solution removed. About 3 to 5 thickeners (not shown) arerequired with a wash ratio (water to solids) of about 5 to 7 to reduceentrained liquor down to less than 100 ppm Cu in the final residue.

The thickener underflow from the last thickener is the final residuestream 35 at about 50% solids. This can be treated for the recovery ofprecious metals, such as gold and silver, or sent to tailings. Therecovery of precious metals will be described later on with reference toFIG. 5.

The main constituents of the stream 35 are hematite and elementalsulphur, which may be recovered by flotation if market conditionswarrant.

The thickener overflow from the first thickener is the product solution33 which is fed to the primary solvent extraction stage 16, as shown. Asan example, this solution contains about 12 g/L Cu, 1 g/L Cl and 0.5 g/LFe.

The optimum copper concentration is determined by the ability of thesolvent extraction stage 16 to extract the maximum copper from thesolution 33. Since a fraction of about one-third of the raffinate fromthe solvent extraction stage 16 is eventually neutralized, it isimportant to minimize the copper content of this raffinate.

Solvent extraction performs best on dilute copper solutions due to thefact that a concentrated copper solution results in a higher acidconcentration in the raffinate which tends to lower extractionefficiency. More concentrated solutions are, however, cheaper to treatfrom a capital cost point of view, since the volume is less. Above acertain point, though, the increased concentration does not reduce thesize of the solvent extraction unit, since (i) there is a maximumorganic loading and (ii) aqueous volume is generally kept equal toorganic volume for mixing purposes by means of aqueous recycle.Therefore, the total volume of organic extractant and aqueous solutionis only determined by the volume of organic extractant. The maximumorganic loading and hence volume of organic is determined by theconcentration and characteristics of the particular organic solventselected. For the typical solvent, e.g. an hydroxy-oxime, the maximumloading per pass at 40% volume concentration in diluent is about 12 g/LCu. Therefore, the product solution 33 also should contain about 12 g/LCu.

The copper is extracted from the product solution 33 from the CCDthickener overflow in two stages of extraction in the primary solventextraction stage 16 to produce a raffinate 37 with about 20 g/L freeacid and about 0.3 to 1 g/L Cu. Most of this raffinate 37 is recycled tothe atmospheric leaching stage 14 but about 25 to 30% is surplus to theacid requirements of the atmospheric leaching stage 14 and must beneutralized. This surplus 121 is split off as shown at 36 andneutralized.

The neutralization is effected in two stages to maximize copper recoveryand to prevent possible environmental problems with the neutralizationresidue due to copper content, i.e. the unrecovered copper from theraffinate 37 will precipitate upon neutralization and can thenre-dissolve later, in a tailing pond, for example.

The first stage neutralization takes place at pH 2 to pH 3, as shown at38, using limerock, which is very economical as a reagent, compared withlime. The neutralization product is filtered at 40 and the resultantsolids are washed with water from the external source 45. The solids,which are mainly gypsum and iron hydroxides, are discarded, as shown at41.

The filtrate 39 is sent to the secondary solvent extraction stage 18 forthe recovery of residual copper values. The secondary solvent extraction18 benefits from the primary neutralization 38 and results in a very lowcopper concentration in the secondary raffinate 43, typically from about0.03 to 0.06 g/L Cu.

As indicated by the broken lines in FIG. 1, the secondary solventextraction stage 18 uses the same organic extractant as the primarysolvent extraction circuit 16. This is also tied in with the solventextraction 27 of the pressure oxidation filtrate bleed 26. The organicextractant which is washed at 42 with wash water 122 from an externalsource 45, and stripped at 44 is recycled to the secondary solventextraction stage 18 and then passes to the primary extraction stage 16.The stripped organic 125 is split to pass a portion thereof to thesolvent extraction 27. The raffinate from the solvent extraction 27 isadded to the loaded organic 123 from the solvent extraction 16 prior tothe wash 42. The wash water 47 from the wash 42 is passed to thepressure oxidation filter 24, to serve as a feed wash water onto thefilter 24. The resultant wash filtrate is added to the pressureoxidation filtrate 29, thus recovering the copper and chloride contentfrom the solvent extraction wash water 47.

The raffinate 43 from the secondary solvent extraction stage 18 isneutralized again in a secondary neutralization stage 46, this time atpH 10 and filtered at 48 to remove all dissolved heavy metals, producinga solution 51 which is used as wash water in the CCD circuit 34 forwashing the final leach residue 35. The solid residue from thefiltration 48 is discarded, as shown at 53.

Stripping of the loaded and washed organic at 44 is effected by means ofspent acid or electrolyte 55 from the electrowinning stage 20 to obtaina pure copper sulphate solution or pregnant electrolyte 57 which is thenpassed to the electrowinning stage 20 for electrowinning in the usualway.

It can be seen that all solution streams in the process are thusrecycled and there are no solution effluents from the process. Onlysolid residues are discarded from the process.

Process Mode B

FIG. 2 is a flow diagram of Mode B. The same reference numerals are usedto indicate stages or steps in the process which correspond with thosein the previous embodiment of FIG. 1. For example, the pressureoxidation stage is again indicated by 12, the atmospheric leach stage by14, the electrowinning stage by 20, the flash tank(s) by 22, thepressure oxidation filtration by 24, the bleed treatment of the pressureoxidation filtrate 29 by reference numeral 28, the grinding stage byreference numeral 30 and the CCD wash circuit by reference numeral 34.

In this mode of the process, the pressure oxidation 12 is carried outboth to oxidize and to leach into solution most of the copper containedin the feed concentrate. Typically about 85-90% of the copper is leachedinto the solution, with only about 10-15% being left in the residue asthe basic copper sulphate.

The conditions of the pressure oxidation stage 12 in the autoclave aresimilar to those in Process Mode A except that the percentage solids islower, i.e. 150-225 g/L.

In this mode of the process, Δ Cu²⁺ ! is typically 30 to 40 g/L Cu, i.e.the copper concentration is greater in the product solution 21 from thepressure oxidation stage 12. The feed solution 25 to the pressureoxidation stage 12 typically contains 10-15 g/L Cu and 12 g/L Cl,together with about 20 to 30 g/L sulphuric acid.

In this mode, no sulphuric acid is added to the pressure oxidation stage12 from an external source, as is the case with the FIG. 1 embodiment.In this mode, the acid is obtained from recycle in the process, i.e. bythe recycle of the pressure oxidation filtrate 29. The product solution21 from the pressure oxidation stage 12 contains about 40 to 50 g/L Cuand 11 to 12 g/L Cl at about pH 2 to 2.5.

The copper leached into the product liquor 21 from pressure oxidationstage 12 must be controlled so as to obtain the desired distribution ofcopper between liquor (85 to 90%) and residue (10 to 15%). Thisdistribution results in a small but important amount of basic coppersulphate solids in the leach residue. The pH is a convenient indicatorof the presence of basic copper sulphate, since it is a buffering agent.With strong copper sulphate concentration in solution, a pH range of 2to 2.5 indicates basic copper sulphate. Below pH 2 almost all the basiccopper sulphate will be dissolved, whereas above pH 2.5, too much basiccopper sulphate is formed and insufficient copper is likely to be foundin the solution 21.

The primary method of control is the amount of acid in the feed liquor25 to the pressure oxidation stage 12. The acid level in turn iscontrolled by the degree of neutralization of the raffinate from solventextraction of the pressure oxidation filtrate 29 raffinate describedbelow. Usually, about 25 to 50% of the acid must be neutralized,depending on the amount of acid that is required.

The acid generated during the pressure oxidation stage 12 varies fromone concentrate to another and according to conditions employed. If theconcentrate produces a large amount of acid during the pressureoxidation stage 12, then the feed solution 25 will need less acid toachieve the desired result. The minimum copper (from concentrate feed)that should go to liquor 21 is about 10%. Below 10%, the pH dropssufficiently low so that iron concentrations in the pressure oxidationfiltrate 29 increase rapidly. Normally, iron is about 10 to 50 ppm, butif pH is below 2 and basic copper sulphate in residue disappears, theniron can increase to above 1 g/L fairly quickly. This is undesirablebecause there are several impurity elements such as As and Sb which areonly removed from solution simultaneously with iron hydrolysis.Therefore, absence of iron in solution is a good guarantee of lowimpurity content in the pressure oxidation filtrate 29. Iron is also animpurity itself that must be avoided in the electrowinning circuit 20 asfar as possible.

There is another factor, however, which places a maximum on Cu insolution. It has been found surprisingly that certain concentratesactually leach more completely if the copper concentration is lower.This is believed to be due to either formation of secondary CuS, asdescribed above, or to some other phenomenon related to poor oxidationcharacteristics of the primary mineral, chalcopyrite, in high copperconcentration solutions. It is found that elemental sulphur, producedduring the reaction in the pressure oxidation stage 12, can coat oractually encapsulate unreacted chalcopyrite particles and hinder theaccess of reagents. This results in poor copper recovery. The phenomenonis apparently accentuated by high Cu levels in solution. It can beovercome or mitigated by the use of surfactants, as described above. Theproblem is more severe with some concentrates, particularly high grade,than others. Therefore, for these concentrates it is desirable to limitthe copper concentration in the pressure oxidation filtrate (i.e.greater than about 95%) over all. To do this, it is necessary to have asubstantial proportion of the copper as basic copper sulphate, i.e. insolid residue from the pressure oxidation stage 12 rather than thepressure oxidation filtrate. Typically, 20-40% of copper may report tosolids, if necessary, to keep the copper concentration low enough toobtain high copper recovery.

Higher grade concentrates exhibit the problem of low copper recoverywith high copper in solution. Therefore, an increasing proportion ofcopper must report to solids as the grade increases. Tests with threedifferent concentrates illustrate this relationship:

    ______________________________________             H.sup.+ /Cu                    Cu Distribution %                                   Total    Conc. #           % Cu    Molar    PO liquor                                   PO residue                                           recovered    ______________________________________    1      41      0.55     0      100     97.3    2      28      0.70     63     37      95.7    3      22      0.96     85     15      94.7    ______________________________________

The H⁺ /Cu molar ratio refers to H⁺ in the feed acid and Cu in the feedconcentrate. The H⁺ in the feed acid is taken to be all the protonsavailable on complete dissociation of the acid even if under existingconditions the acid is not completely dissociated. The H⁺ shown in thetable is optimum level found by experiment to give the best results.

For concentrate #1, which was a high grade concentrate, the processchosen is Mode A, where all of the copper reports to the leach liquor 33and Δ Cu²⁺ !=0. The H⁺ /Cu ratio is that found which was necessary byexperimentation to give the desired result of Δ Cu²⁺ !=0.

For concentrate #2, a medium grade concentrate, Mode B was chosen, butwith a substantial amount of the copper reporting to the solid basiccopper sulphate. This was achieved by keeping the H⁺ /Cu ratio lowenough so that not all of the copper dissolved into the liquor.

For concentrate #3, a low grade concentrate, Mode B was also chosen butin this case, the minimum amount of copper reported to the residue, byadjusting the H⁺ /Cu ratio to be high enough.

The residue from the pressure oxidation stage 12 is leached 14 withraffinate 37 returning from the solvent extraction 16 which is diluteacid, at 3-10 g/L H₂ SO₄. Since most of the copper from the pressureoxidation stage 12 reports to the pressure oxidation filtrate 29 andonly a small fraction of the pressure oxidation residue, the resultantleach liquor 31 from the atmospheric leach 14 is quite dilute in copper.In turn, this produces a dilute raffinate 37 from the solvent extraction16. Typically, the atmospheric leach liquor 31 is 3-7 g/L Cu and 0.2 to0.5 g/L Fe.

The slurry resulting from the atmospheric leaching stage 14 is difficultto filter, as was the case with Mode A. Good liquid/solid separation andwashing, however, can be achieved as before using a series of thickenersin a CCD arrangement 34. Wash water 51 is provided by raffinate from thesolvent extraction 16, which is neutralized, as indicated at 46. This issimilar as in Mode A. The only major difference is the lower tenor ofthe solution 33 and the reduced volume.

The solution 33 produced by the atmospheric leaching stage 14 issubjected to the solvent extraction 16. The copper containing solution29 from the pressure oxidation stage 12, is subject to a solventextraction stage 50. There are, therefore, two solvent extractionoperations, i.e. 16 and 50, treating two different streams of liquor 33and 29, respectively. It is a feature of the process according to theinvention that the organic extractant used for effecting the solventextraction operations is common to both solvent extractions 16 and 50.

As shown in FIG. 2, the stripped organic 125 coming from the commonstripping operation 44 is first introduced into the solvent extractioncircuit 16, which has the weakest copper concentration in the aqueousfeed stream 33 and therefore needs the organic extractant to be as lowas possible in loading to be effective.

The loaded organic 126 from solvent extraction 16 is then sent to thesolvent extraction 50 where it contacts the higher copper concentrationliquor 29. It is not necessary for the solvent extraction 50 to achievea high extraction ratio because the raffinate 63 from this extraction isrecycled to the pressure oxidation stage 12, as shown. On the otherhand, the raffinate 37 from the solvent extraction 16 is only partlyrecycled and part is neutralized 46 to remove excess acid from thecircuit. Therefore, it is more important to achieve high copper recoveryfrom the solvent extraction 16.

The raffinate 37 from the solvent extraction 16 is split at 36 as inMode A, with about one-third 121 to the neutralization 46 and two-thirds120 recycled to the atmospheric leach stage 14. An important differencefrom Mode A is that the raffinate 37 from solvent extraction 16 issufficiently low in copper, i.e. below 100 ppm, so that it is notnecessary to have a secondary solvent extraction stage beforeneutralization 46, as was the case in Mode A. This is due to the lowercopper concentration and solution volume, allowing the solventextraction 16 to be more efficient.

The loaded organic 65 produced by the two solvent extraction operations16, 50 in series, is washed in two stages in counter current fashionwith dilute acidic aqueous solution 122, as shown at 42. This isprimarily to remove entrained aqueous solution from the loaded organic65 and in particular to reduce the chloride content before the organicgoes to stripping at 44. The amount of wash water required is about 1-3%of the organic volume. The resultant wash liquor 47 produced is recycledto the pressure oxidation stage 12.

The washed organic 69 is stripped at 44 with spent electrolyte 55 fromthe electrowinning stage 20 to provide a pure copper solution orpregnant electrolyte 57 for electrowinning in the usual way.

The raffinate 63 is split at 70 in two portions 72, 74 as determined bythe required molar ratio of H⁺ /Cu. The portion 72 is recycled to thepressure oxidation stage 12. The portion 74 is neutralized at pH 2 withlimerock at 76 and filtered 78. The solid residue is washed anddiscarded, as shown at 80. The filtrate 82 is recycled with the portion72 to form the feed solution 25 to the pressure oxidation stage 12.

A novel feature of the process, therefore, is the use of a commonorganic to extract copper from two separate aqueous feed liquors. Thisprovides considerable economies in lower capital and operating costs inthe solvent extraction circuits. Also, it allows for the use of copiousamounts of water in the atmospheric leaching CCD circuit, so that goodwashing can be achieved on the final residue and yet still recovercopper from such a dilute liquor.

It has been found that the degree of sulphur oxidation that occurs inthe pressure oxidation stage 12 is highly dependent on the type ofconcentrate, such as grade and mineralogy of the concentrate beingtreated, as well as the conditions of the pressure oxidation stage 12.Certain concentrates exhibit considerably higher sulphur oxidation, e.g.25-30%, i.e. oxidation of the sulphur in the concentrate to sulphate,and the effect is particularly marked with the low grade concentrateswith less than about 28% Cu by weight. The inventor has found that thesignificance of this variation is not so much the copper grade itselfbut the copper/sulphur ratio in the concentrate. The main impurityelements in a copper concentrate are iron and sulphur due to the factthat copper ores are generally composed of chalcopyrite together withother minerals, particularly pyrite FeS₂ or pyrrholite FeS.

Process Mode B deals with the problem of excess sulpfur oxidation in thepressure oxidation stage 12 when lower grade concentrates are used bydeliberately dissolving 90% of the copper and minimizing the formationof basic copper sulphate. The reaction for chalcopyrite is:

    CuFeS.sub.2 +5/4O.sub.2 +H.sub.2 SO.sub.4 →CuSO.sub.4 +1/2Fe.sub.2 O.sub.3 +2S.sup.0 +H.sub.2 O                              (9)

The filtrate 29 from the pressure oxidation stage 12 thus contains highlevels of copper sulphate and copper chloride and this is treated in thesolvent extraction stage 50 to produce a pure copper sulphate solutionwhich goes to the electrowinning stage 20.

It has been found that there is a limit to the amount of sulphuroxidation that can be accommodated by the process Mode B. If the sulphuroxidation is high enough and sufficient acid is generated duringpressure oxidation, there will be a surplus of acid left over afterpressure oxidation, even if no acid is added to the feed, such as in theform of acidic raffinate. In this situation, not only will all thecopper in the concentrate be converted to dissolved copper sulphate, butalso some of the iron in the concentrate will be solubilized by thesurplus acid, e.g. as ferric sulphate.

It is desirable that iron in the concentrate report to the pressureoxidation residue as stable hematite, Fe₂ O₃, and not to the solution,where it must be separated from the copper. Typical concentrates have anFe:Cu ratio of at least 1:1, and therefore the efficient and completeelimination of Fe at an early stage is an important aspect of theprocess. Other impurities such as arsenic, antimony, etc., are alsoremoved with iron by co-adsorption or precipitation mechanisms.

It has been found that some concentrates, however, exhibit so muchsulphur oxidation (acid generation) that the acid-consuming capacity ofpressure oxidation is exceeded, and some iron is leached into solution,even under process Mode B conditions. It is a target of the process toproduce a low-iron liquor, typically with 0.05 g/L Fe. Some concentrateswhich have been tested have produced pressure oxidation liquors with 1.0to 12.0 g/L Fe. Similarly the pH of the pressure oxidation liquor isnormally targeted to be in the range 2.0 to 3.5, corresponding to lessthan 1 g/L free acid, but concentrates tested have produced pressureoxidation liquors with pH in the range 1.2-2.0, corresponding to 1 to 15g/L free acid.

Accordingly, a further embodiment of the process has been developed,i.e. process Mode C for the treatment of the above concentrates, termed"Mode C" concentrates. Process Mode C will now be described below.

Process Mode C

The Mode C concentrates that exhibit a strong tendency towards sulphuroxidation and hence acid generation are those with a high S:Cu ratio, ormore generally S:M ratio, where M=base metals, such as Cu, Zn, Ni, Co,Pb, etc., but not including Fe, which does not consume acid.

Nickel or nickel/copper concentrates may often be Mode C, because theyare frequently low-grade, with S:M ratio often about 2:1 or higher. Somecopper or copper/gold concentrates are also Mode C, if they are lowgrade because of high pyrite content. Some copper/zinc concentrates havealso been found to be high in pyrite and hence of Mode C type as well.

In general there is a correlation between pyrite (FeS₂) content and thetendency toward Mode C type behaviour. However, there are alsoexceptions to this trend, as not all pyrites react in the same way. Somepyrites oxidize sulphur more readily than others. In contrast,pyrrhotite (Fe₇ S₈) or the related iron-zinc mineral sphalerite,(Zn,Fe)S, appear to result in much less sulphur oxidation, and thusexhibit Process Mode A or B behaviour.

In this process mode the objective, as with Modes A and B, is to reducesulphur oxidation during pressure oxidation by the addition of sulphateor sulphuric acid. The pressure oxidation filtrate resulting from thehigh acid concentration will have high levels of dissolved iron and lowlevels of dissolved copper. Therefore, a neutralizing agent, such aslime or limestone, is added to neutralize the pressure oxidation slurryprior to filtration.

Process Mode C is essentially a special case of Process Mode B, with twoimportant differences.

First, all the raffinate 63 (FIG. 2) is neutralized, before returningthis stream to the pressure oxidation 12, i.e. there is no raffinatesplit, with one part being neutralized and the other part by-passing theneutralization.

Secondly, the pressure oxidation slurry before filtration 24 of theleach residue is subjected to an extra neutralization step, the"pressure oxidation neutralization", to neutralize excess acid andprecipitate any Fe in solution at this time.

A convenient opportunity for the pressure oxidation neutralization is ina conditioning tank after flash let-down 22 to atmospheric pressure,when the slurry is at or near the boiling point of the solution, i.e.about 80°-95° C.

However, there is an inherent problem with this, i.e. the unwantedformation of gypsum deposits in subsequent unit operations downstreamfrom the neutralization, where the temperatures are lower. The gypsumdeposits because the solubility of calcium sulphate is lower, causingthe supersaturation of calcium sulphate once the temperature is reduced.

If the above-described neutralization is carried out on a slurry at80°-95° C., using limerock, then the resulting solution will besaturated with calcium sulphate at that temperature. If the resultingsolution is subsequently cooled to 40°-50° C. for solvent extraction,then the solubility of calcium sulphate is markedly reduced andconsequently there will be slow precipitation of solid calcium sulphate,most likely the dihydrate form known as gypsum, CaSO₄.2H₂ O. Such gypsumis well known to form tenacious scale in piping, valves, tanks, etc. andcause severe operational problems in a commercial plant.

This problem has been solved by carrying out the neutralization insidethe autoclave or pressure vessel 300 (FIG. 3), at the conclusion of thepressure oxidation stage 12, when the temperature is about 115° to 160°C. It has been found that in this temperature range, the solubility ofcalcium sulphate is equal to or lower than at the lower temperatures,where solvent extraction takes place. Thus the saturation level forcalcium sulphate produced during the neutralization is equal to or lowerthan at any subsequent stage in the process, and supersaturation doesnot occur due to a temperature drop. In this way, gypsum scalingproblems are avoided.

The pressure vessel 300 in the present example has five compartments302.

To achieve the neutralization inside the pressure vessel 300, it hasbeen found preferable to use slaked lime slurry, Ca(OH)₂, rather thanlimerock, CaCO₃, as the active neutralizing agent as indicated at 304 inFIG. 3. Slaked lime avoids the formation of carbon dioxide gas, CO₂,which is attendant upon the reaction of limerock with acid. CO₂ gasoccupies a large volume of space within the pressure vessel 300,otherwise needed for oxygen, and effectively shuts down the desiredpressure oxidation reaction. To use slaked lime in a continuous reactionvessel, it is necessary to pump it into the pressure vessel 300 in thecustomary slurry form at 10-20% solids in water, into the lattercompartments 302 of the vessel 300. Thus it is beneficial to carry outthe pressure oxidation in the first three or four compartments 302, andpump in the slaked lime slurry to the last or last but one compartment302. The concentrate, H₂ SO₄, chloride and oxygen are introduced intothe autoclave 300, as indicated at 306, 307 and 308, respectively.

The amount of slaked lime to be used is determined by the amount of acidand iron to be neutralized and the amount of copper that may be neededto be precipitated in the form of basic copper sulphate. Generally it isdesirable to finish pressure oxidation 12 with no free acid andvirtually no iron in solution, i.e. less than 10 ppm Fe, and a pH ofabout 2.5 to 4.0.

As already indicated above, it is important to keep the addition ofwater to the system to a minimum. This also applies to the pressureoxidation neutralization using Ca(OH)₂ (slaked lime). Normally a solid'scontent of about 10-20% is the maximum that can be tolerated before theviscosity of the slaked lime slurry becomes too difficult to handle.This is particularly a problem when Ni is present in the concentratewhere the consumption of Ca(OH)₂ is high. This problem can be overcomeby the addition of a viscosity modifier, such as caustic, potash orlignosol. This effectively reduces the viscosity so that a 30% or highersolids content can be tolerated.

The resultant slurry (indicated at 309), now at pH 2.5 to 4.0 is flashedin two stages to atmospheric pressure and then filtered (24) (FIG. 2).The filter cake is washed to remove entrained liquor (Cu,Cl) as much aspractical. The filter cake now containing solid calcium sulphate,produced at the pressure oxidation temperature, together with othersolids, such as hematite, elemental sulphur and basic copper sulphateproceeds to atmospheric leaching 14 where any precipitated copper isleached as usual at about pH 1.5-1.8, and the resultant residue washedthoroughly in the CCD circuit 34. The filtrate 29 from the pressureoxidation filtration is treated as in process Mode B for Cu removal bythe solvent extraction stage 50, producing a raffinate 63 that then goesto neutralization 76, as before, and then recycled back to the pressureoxidation 12, but without the raffinate split 70, as indicated above.

Although the pressure oxidation stage 12 is chloride catalyzed, it doesnot use a strong chloride solution, e.g., in the preferred embodimentonly about 12 g/L is needed which will support about 11 g/L Cu or Zn asthe respective chloride salt. If a higher concentration of metals isneeded or produced, it is as the sulphate salt. Thus, the solutionsproduced by the pressure oxidation stage 12 are generally mixtures ofthe sulphate and chloride salts, not pure chlorides.

Tests were carried out on a low grade sulphide ore to investigate theeffect of the pressure oxidation neutralization.

In a first test, which was operated according to Mode B with noneutralization, the feed to the autoclave comprised 10.7 g/l free acid,12 g/l Cu and 12.5 g/l chloride in solution. The resultant pressureoxidation filtrate 29 after filtration was at a pH of 1.72 with a copperconcentration of 48 g/l and dissolved iron present in an amount of 2350ppm. The solid residue sent to the atmospheric leach 14 contained 2.0%Cu.

In a second test, on the same low grade concentrate, which was operatedaccording to Mode C with a neutralization step in the autoclave, thefeed to the autoclave comprised 16.0 g/l free acid, 14 g/l Cu and 12 g/lchloride in solution. The resultant pressure oxidation filtrate 29 wasat a pH of 3.05 with a copper concentration of 42 g/l and only 25 ppmiron in solution. The solid residue contained 6.5% Cu. No problem withgypsum precipitation later on in the lower temperature stages of theprocess was encountered.

In both tests the sulphur oxidation, i.e. oxidation of the sulphur inthe concentrate to sulphate, was about 27-30%.

These tests illustrate that it is possible to control pH byneutralization in the autoclave, thereby minimizing the iron content inthe pressure oxidation filtrate, without the problem of gypsumprecipitation in the system.

Referring to FIG. 4, an embodiment of the process which is suitable forthe treatment of copper-zinc concentrates with about 20-25% Cu and about1-10% Zn is shown. The process is similar to the process Mode B of FIG.2 and like reference numerals are again used to indicate correspondingsteps. FIG. 4 is less detailed than FIG. 2.

It has been found that good extraction of Zn in the pressure oxidationstage 12 can be achieved if sufficient acid is added to the feedsolution to maintain the final pH of the slurry below about pH 2.Otherwise, the conditions are similar as for Cu concentrates beingtreated by the FIG. 2 process, i.e., 150° C., 200 psig (1400 kPa) O₂, 12g/L Cl.

In the Mode B type process, Cu is primarily solubilized during thepressure oxidation stage 12, and is extracted by the Cu solventextraction 50. This solvent extraction stage 50 is operated inconjunction with the Cu solvent extraction stage 16 in which Cu isextracted from the leach liquor coming from atmospheric leach stage 14,as described with reference to FIG. 2 above. The solvent extractions 50,16 produce a concentrated copper solution which is treated in anelectrowinning stage 20, as described above.

The residue 35 from the atmospheric leach stage 14 is treated forsulphur and precious metals recovery (39) as will be described belowwith reference to FIG. 5.

The raffinate 37 from the Cu solvent extraction 16 is split into twostreams 120 and 121 in the ratio of 2/3 to 1/3, as in the FIG. 2embodiment. The stream 120 is recycled to the atmospheric leach 14,whereas the stream 121 is subjected to neutralization 46 at a pH ofabout 4 and then subjected to a liquid/solid separation 48.

The raffinate 63 from the Cu solvent extraction 50 is subjected toneutralization 76 at pH 2 with limerock and then subjected to aliquid/solid separation 78. The solid gypsum residue 80 is discardedwith the gypsum residue 53 from the liquid/solid separation 48.

The liquid phase from the liquid/solid separation 78 is subjected, withthe liquid phase from the liquid solid separation 48, to solventextraction 246 with a suitable zinc extractant, such asdiethylhexylphosphoric acid (DEHPA), to produce a loaded Zn organic.This organic stream is carefully purified of Cu, Co, Cd, Cl, etc.,before stripping with spent acid from a subsequent zinc electrowinningstage. The purification can be effected by scrubbing of the loadedorganic using aqueous ZnSO₄ solution. The raffinate is recycled to thepressure oxidation stage 12.

The raffinate from the Zn solvent extraction is recycled to the pressureoxidation stage 12. Any traces of DEHPA left in the raffinate will besubjected to the highly oxidizing conditions of the pressure oxidationstage 12 to counteract contamination of the hydroxy-oxime copperextractant with DEHPA. It has been found that contamination of thehydroxy-oxime reagent with DEHPA results in the deterioration of theformer reagent.

The recovery of precious metals, such as gold and silver, will now bedescribed with reference to FIGS. 5A and B. This process involves thetreatment of the final residue stream 35 in FIGS. 1, 2 and 4.

The precious metals are not leached during the pressure oxidation stage12 but remain in the solid residue 35 remaining after the atmosphericleaching stage 14.

In order to facilitate precious metal recovery, the flash down 22 fromthe pressure oxidation stage 12 is carried out in two stages. The firststage is at a temperature slightly above the freezing point of elementalsulphur, i.e. about 120° to 130° C. with a corresponding steam pressureof about 10 to 20 psig (70-240 kPa). The process is preferably carriedout in a continuous mode, the retention time at the first flash let-downstage being about 10 to 30 minutes.

The second flash let-down stage is at atmospheric pressure and about 90°to 100° C. with a retention time of again at least 10 minutes. Thisallows the elemental sulphur, which is still molten in the firstflash-down stage, to convert to one of the solid phases, such as thestable orthorombic crystalline phase. This procedure facilitates theproduction of clean crystals of elemental sulphur which is important tothe recovery of the precious metals from the leach residue.

The leach residue 35 now produced by the atmospheric leaching stage 14contains, in addition to the precious metals, hematite, crystallineelemental sulphur, unreacted sulphides (pyrite) and any additionalproducts that may result from the particular concentrate being used,e.g. gypsum and iron hydroxides.

Gold in the residue 35 is believed to be largely untouched by theprocess so far and most likely is in the native state. Silver, however,is oxidized in the pressure oxidation stage 12 and is probably presentas a silver salt, such as silver chloride or silver sulphate.

It has been found that conventional cyanidation does not leach gold wellfrom the residue 35. It is believed that this is due to theencapsulation of the gold in mineral particles, such as pyrite. The goldcan however be liberated by the pressure oxidation of these minerals,referred to as "total oxidative leaching" or "pyrite leach". In order toeffect such leaching without oxidizing elemental sulphur also containedin the residue 35, the process comprises the step of removing as much ofthe elemental sulphur as possible.

Firstly, by virtue of the two stage flash-down, good quality sulphurcrystals are produced. Secondly, the leach residue 35 is subjected tofroth flotation 402 to produce a sulphur rich flotation concentrate 404and a sulphur depleted flotation tail 406. The tail 406 is subjected toa solid/liquid separation 408 to produce a liquid which is recirculatedto a conditioning tank 410 upstream of the flotation step 402 and asolid 412 which is sent to the total oxidative leaching stage 414.

The flotation concentrate 404 is filtered (416), dried to a low moistureand then melted in a melting step 418 at about 130° C.-150° C. toproduce a slurry 420 of liquid sulphur and particles of solid minerals.

The slurry 420 is filtered (422) to remove the liquid sulphur which isthen cooled (424) to produce an elemental sulphur product 426. Thecooled sulphur can be subjected to an optional sulphur purification step425 to remove impurities such as selenium and tellurium therefrom.

The solid residue from the filtration 422 is subjected to a hot sulphurextraction step 428 at 90° C. with kerosene or other suitableextractant, such as perchloroethylene. The resulting hot slurry isfiltered (430) to produce a low sulphur (less than 5% elemental sulphur)residue 432 which is sent to the total oxidative leach 414. The hotfiltrate is cooled (434) to reduce the solubility of sulphur, producingcrystalline S° which is filtered off at 436, to return the kerosenewhich is recycled to the sulphur extraction step 428.

A test was carried out in which 100 g of residue from the atmosphericleach 14 containing 25.1% elemental sulphur (S°) and 3% sulphide wasprocessed through flotation 402, melting 418 and extraction 428. Thisproduced 73.8 g of desulphurized residue (feed material for the totaloxidation leach 414) containing 1.9% S° and 4.1% sulphide, i.e. a totalof 6% total sulphur.

The desulphurized residue contained 5.9% of the elemental sulphur (S°)in the original leach residue, i.e. 94.1% was recovered to a pureelemental sulphur product.

The total oxidative leach 414 is carried out at about 200° C.-220° C.and 50-150 psig (500-1200 kPa) oxygen partial pressure, sufficient tofully oxidize all sulphur and metal compounds to the highest valences,respectively. Thus all sulphur and pyrite are oxidized to sulphate. Theoxidation is conducted in acidic conditions, such as with the acid beingproduced in situ. The reaction is highly exothermic and generally thedesired operating temperature can be achieved even with a cold feedslurry, provided there is sufficient fuel present as sulphur in thesolid feed. Typically about 6 to 10% of total sulphur will be sufficientwith normal percentage solids in the feed slurry.

After the total oxidative leaching 414, the slurry is subjected toneutralization 437 at pH 2-3 with limestone and then filtered (438) toproduce a solid residue containing precious metals and a filtrate whichis generally acidic and which may contain base metal values, such ascopper, which can be extracted by an optional solvent extraction step440 and sent to the main solvent extraction circuit. The resultantraffinate is recycled to the total oxidation leach 414, as indicated at442.

Prior the cyanidation 444, the solids from the filtration 438 can besubjected to an optional lime boil step 443 to facilitate the recoveryof silver during the cyanidation 444 by the decomposition of silverjarosite compounds formed during the total oxidative leach 414.

The precious metals are in the solids remaining after the filtration438. Now that pyrite and other encapsulating minerals in the originalconcentrate have been decomposed, the precious metals are amenable tocyanidation 444.

In the cyanidation step 444, the solids are leached with NaCN underalkaline conditions. In order to effect this, the solids are slurried upwith cyanide solution to form a 30-40% solids slurry. Additional NaCNand slaked lime are added as required to maintain a minimum NaCNconcentration of about 0.2 to about 0.5 g/L NaCN, with a pH of about 10.The temperature is ambient and usually about 4 to 8 hours retention timeis required in continuous mode of operation.

Both gold and silver report in high yield to the cyanide solution, andare recovered typically by the established process of carbon-in-pulpcircuit, whereby activated carbon is added to the cyanide slurry toabsorb the precious metals, without the necessity of filtration. Theloaded carbon, now rich in precious metals is separated by screening(445) and the barren pulp discarded to tailing.

The loaded carbon is treated by established methods to recover theprecious metals content by a leach/electrowin/smelt process (447). Theproduct is generally Dore metal containing both gold and silver, whichis sent to a gold refinery 449 for final separation of gold from silver.Barren carbon from a carbon regeneration step 451 after the preciousmetals recovery, is recycled to the carbon-in-pulp circuit 444.

The overall recovery of precious metals by the total process isgenerally well over 90%, and under optimum conditions approach 99%.

A test was carried out in which desulphurized residue was processed in atotal oxidative leach 414 at 220° C. for 2 hours under oxygen pressureand then depressurized and cooled to room temperature. The resultantslurry was neutralized to pH 3 with limestone and then filtered. Thefiltered cake was then leached with cyanide solution under standardconditions to leach gold and silver.

The gold extraction after the total oxidative leach 414 and cyanidation444 was 97% with only 1.0 kg/t NaCN consumption. In comparison, the goldextraction on a residue that had not been oxidized in the totaloxidative leach 414 was only 34% and cyanide consumption was extremelyhigh at 19.0 kg NaCN/t.

What is claimed is:
 1. A process for the extraction of zinc from asulphide ore or concentrate containing copper and zinc, comprising thesteps of:subjecting the ore or concentrate to pressure oxidation in thepresence of oxygen and an acidic halide solution to produce a liquorcontaining copper and zinc in solution and a solid residue; subjectingthe liquor to a first solvent extraction with a copper extractant toremove copper from the solution and to produce a copper depletedraffinate; subjecting the copper depleted raffinate to neutralization toraise the pH thereof; subjecting the resulting copper depleted raffinateto a second solvent extraction with a zinc extractant to produce a zincdepleted raffinate; and recycling the zinc depleted raffinate to thepressure oxidation.
 2. A process according to claim 1, wherein thehalide is selected from chloride or bromide.
 3. A process according toclaim 2, wherein the halide is chloride and the concentration of thechloride ions in the acidic solution is about 12 g/l.
 4. A processaccording to claim 1, wherein the ore or concentrate also containsprecious metals and further comprising the steps of:removing elementalsulphur from the solid residue resulting from the pressure oxidation toobtain a low elemental sulphur residue; and subjecting the low elementalsulphur residue to an oxidative leach at elevated temperature andpressure to oxidize elemental sulphur and sulphide minerals present inthe low sulphur residue to produce a residue for the extraction of theprecious metals therefrom.
 5. A process according to claim 4, whereinthe sulphur removal comprises the steps of:subjecting the solid residueresulting from the pressure oxidation to froth flotation to produce asulphur rich flotation concentrate and a sulphur depleted flotationtail; and subjecting the flotation concentrate to sulphur extractionwith a sulphur extractant to produce the low elemental sulphur residue.6. A process according to claim 5, wherein the sulphur depletedflotation tail is subjected to a solid/liquid separation to produce aliquid which is recirculated to the froth flotation and a solid which issubjected to the oxidative leach.
 7. A process according to claim 5,wherein the sulphur extraction of the flotation concentrate is effectedat a temperature of about 90°-150° C.
 8. A process according to claim 7,wherein the sulphur extractant is selected from the group consisting ofkerosene and perchloroethylene.
 9. A process according to claim 4,wherein the oxidative leach is carried out at a temperature of about200°-220° C. and an oxygen partial pressure of about 500-1200 kPa underacidic conditions.
 10. A process according to claim 4, wherein prior tothe sulphur removal, the pressure oxidation slurry is flashed toatmospheric pressure in a two stage let-down in which the first stage isat a temperature above the freezing point of elemental sulphur.
 11. Aprocess according to claim 10, wherein the first stage let-down is at atemperature of about 120° to 130° C. and a steam pressure of about 10 to20 psig (170 to 240 kPa).
 12. A process according to claim 10, whereinthe second stage let-down is at a temperature of about 90° to 100° C.and at atmospheric pressure.
 13. A process according to claim 1, furthercomprising the step of adding acid to the pressure oxidation to maintainthe resulting pressure oxidation slurry at a pH of below about pH 2.